Methane conversion to higher hydrocarbons

ABSTRACT

The present invention provides a process for the manufacture of acetylene and other higher hydrocarbons from methane feed using a reverse-flow reactor system, wherein the reactor system includes (i) a first reactor and (ii) a second reactor, the first and second reactors oriented in a series relationship with respect to each other, the process comprising supplying each of first and second reactant through separate channels in the first reactor bed of a reverse-flow reactor such that both of the first and second reactants serve to quench the first reactor bed, without the first and second reactants substantially reacting with each other until reaching the core of the reactor system.

CROSS REFERENCE TO RELATED APPLICATIONS

This application is a divisional of U.S. application Ser. No.11/643,541, filed Dec. 21, 2006 now U.S. Pat. No. 7,943,808, whichclaims benefit of and priority to U.S. Provisional Application No.60/753,961, filed Dec. 23, 2005, the entirety of which is incorporatedherein by reference. This application is also a continuation in part ofU.S. Ser. No. 11/639,691, filed Dec. 15, 2006 now U.S. Pat. No.7,846,401 which is not incorporated by reference herein.

FIELD OF THE INVENTION

The invention relates to the manufacture of acetylene from methane. Thepresent invention also relates broadly to regenerative reactors. Moreparticularly the invention relates to an improved process and apparatusfor producing acetylene from a methane feed by controlling combustionfor thermal regeneration of reverse flow regenerative reactors in aunique and thermally efficient way.

BACKGROUND OF THE INVENTION

Acetylene (or ethyne, HC≡CH) has long been recognized as one of the fewcompounds that can be made directly at high selectivity from methane butthe conditions of that manufacture have placed it beyond commercialpracticality for other than high cost, specialty production. Acetylenecan be converted to a number of other desirable hydrocarbon products,such as olefins and vinyls. One of the biggest impediments to producingacetylene from methane feeds has been the very high temperaturesrequired to produce high-yield conversion of methane to acetylene. Manyof the desired products that could be manufactured from the producedacetylene are today instead being produced via more economicalprocesses, such as thermal cracking of higher molecular weighthydrocarbon feeds such as ethane and naphthas, in thermal crackers. Thehigher molecular weight feed crack at lower temperatures than methane.Equipment, materials, and processes were not previously identified thatcould continuously withstand the high (>1600° C.) temperatures requiredfor methane pyrolysis. Pyrolyzing large quantities of methane had beenconsidered much too costly and impractical due to the special types andcosts of equipment that would be required. The developed processes forproducing acetylene have all operated commercially at lower temperaturesfor steam cracking of higher weight hydrocarbon feeds.

It is known that acetylene may be manufactured from methane in smallamounts or batches, using a high temperature, short contact time,yielding a mixture of acetylene, CO, and H₂. Comprehensive discussionsare provided in the Stanford Research Institute report entitled“Acetylene,” a Process Economics Program, Report No. 16, September 1966,and in the Fuel Processing Technology publication (42), entitled“Pyrolysis of Natural Gas: Chemistry and Process Concepts,” by Holmen,et. al., 1995, pgs. 249-267. However, the known processes areinefficient, do not scale well, and are generally only useful forspecialty applications.

The known art discloses that to efficiently obtain relatively highyields of acetylene, such as in excess of 50 wt % or more preferably inexcess of 75 wt % acetylene from the methane feed, temperatures arerequired to be in excess of 1500° C. and preferably in excess of 1600°C., and with short contact times (generally <0.1 seconds) to preventbreaking the acetylene into elemental carbon and hydrogen components.Such temperature and processes have largely been unattractive due to thedegradation of the equipment utilized. Virtually any metal componentsthat are exposed to such temperatures will be costly and willunacceptably degrade.

In addition to the above references, U.S. Pat. No. 2,813,919 disclosesacetylene manufacture from methane in a reverse-flow reactor (aregenerative furnace), operating at temperatures of typically 2500° F.(1370° C.), but up to 3000° F. (1650° C.). U.S. Pat. No. 2,885,455discloses a reverse-flow reactor (a regenerative pebble-bed reactor) forproduction of acetylene from light hydrocarbons. Ethane and propanefeeds are discussed and claimed; methane is not mentioned. Reactiontemperatures up to 3000° F. (1650° C.) and contact times of 0.1 secondor less are disclosed. U.S. Pat. No. 2,886,615 describes a reverse-flowreactor (a regenerative pebble-bed reactor) useful for processinghydrocarbon feed stocks (including natural gas) with hydrogen reactantto prepare olefins, acetylenes, and other product. Temperatures inexcess of 3000° F. (1650° C.) and reaction times of 0.001 to 1 secondare disclosed. The improvement taught is a secondary heat reservoir.

U.S. Pat. No. 2,920,123 describes pyrolysis of methane to produceacetylene at temperatures of 2820° F. (1550° C.) to 2910° F. (1600° C.)and contact time in the range of 0.004 to 0.015 seconds. The exemplifiedreactor is an electrically heated ceramic tube, and the soft carbonproduced as a byproduct under these conditions is removed by oxidationafter 5 seconds of feed.

U.S. Pat. No. 3,093,697 discloses a process for making acetylene byheating a mixture of hydrogen and a hydrocarbon stock (e.g., methane) ata reaction temperature that is dependent upon the particular hydrocarbonemployed, for about 0.01 to 0.05 second. The reference indicates that areaction temperature of 2700° F. to 2800° F. (about 1482° C. to about1538° C.) is preferred for methane and that lower temperatures arepreferred for higher molecular weight hydrocarbons.

U.S. Pat. No. 3,156,733 discloses a process for the pyrolysis of methaneto acetylene and hydrogen. The process involves heating amethane-containing stream in a pyrolytic reaction zone at a maximumtemperature above 2730° F. (1500° C.) and sequentially withdrawing agaseous product from the reaction zone and quenching the product rapidlyto a temperature of about 1100° F. (600° C.) or less. U.S. Pat. No.4,176,045 discloses a process for the production of olefins bysteam-cracking normally liquid hydrocarbons in a tubular reactor whereinthe residence time in the tubes is from about 0.02 to about 0.2 secondand the formation of coke deposits in the tubular reactor is minimized.U.S. Pat. No. 4,929,789 discloses a process for pyrolyzing or thermalcracking a gaseous or vaporized hydrocarbon feedstock using a novelgas-solids contacting device and an oxidation catalyst. U.S. Pat. No.4,973,777 discloses a process for thermally converting methane intohydrocarbons with higher molecular weights using a circulating methaneatmosphere in a ceramic reaction zone.

Chemical Economy and Engineering Review, July/August 1985, Vol. 17, No.7.8 (No. 190), pp. 47-48, discloses that furnaces have been developedcommercially for steam cracking a wide range of liquid hydrocarbonfeedstocks using process reaction times in the range of 0.05 to 0.1second. This publication indicates that the use of these furnacespermits substantial increases in the yield of olefins (i.e., ethylene,propylene, butadiene) while decreasing production of less-desirableco-products. GB 1064447 describes a process for production of acetylenefrom pyrolysis of methane and hydrogen (1:1 to 39:1 H₂:CH₄; clm. 9) inan electrically heated reactor, and quenching with a dry, oxygen-freegas stream. The maximum temperature is 1450 to 2000° C. (preferably 1450to 1750° C.; clm. 2).

The “Wulff” process represents one of the more preferred commercialprocesses for generation of acetylene. The Wulff process includes areverse-flow thermal pyrolysis process and began development in the1920's. Various related processes operated commercially up to about the1960s. These processes typically used feeds heavier than methane andthereby operated at temperatures of less than 1500° C. The most completedescription of the Wulff process is provided in the Stanford ResearchInstitute's “Acetylene”, Process Economics Program Report Number 16(1966). Among the relevant patents listed in this report are U.S. Pat.Nos. 2,319,679; 2,678,339; 2,692,819; and 3,093,697, discussed above. Itis believed that all commercial acetylene plants operated on feeds ofethane, naphtha, and/or butane, but that none have successfully operatedon methane feeds. Wulff discloses a cyclic, regenerative furnace,preferably including stacks of Hasche tiles (see U.S. Pat. No.2,319,579) as the heat exchange medium. However, to contain the locationof the reaction heat generated by the exothermic combustion process, oneof either the fuel or oxygen is introduced laterally or separately intothe central core of the reactor where it mixes with the other reactioncomponent. The other reaction component is preferably introduced throughthe reactor tiles to cool the reactor quench section. Thereby,combustion may occur in a known location within the reactor. However,this also exposes the lateral injection nozzles or ports to thecombustion product, including the extremely high temperature needed tocrack methane feeds. Degradation in nozzle performance, shape, and/orsize consequently made it extremely difficult to control flame shape,temperature, and efficiency. Although some of the Wulff art disclose useof various refractory materials, a commercially useful process formethane cracking was not achieved utilizing such materials. Also, afurther drawback of the Wulff process is that the laterally orseparately introduced portion of exothermic reactant is not utilized forquenching the recuperation reactor bed.

Regenerative reactors, including reactors such as disclosed by Wulff,are typically used to execute cyclic, batch-generation, high temperaturechemistry. Typically, regenerative reactor cycles are either symmetric(same chemistry or reaction in both directions) or asymmetric (chemistryor reaction changes with step in cycle). Symmetric cycles are typicallyused for relatively mild exothermic chemistry, examples beingregenerative thermal oxidation (“RTO”) and autothermal reforming(“ATR”). Asymmetric cycles are typically used to execute endothermicchemistry, and the desired endothermic chemistry is paired with adifferent chemistry that is exothermic (typically combustion) to provideheat of reaction for the endothermic reaction. Examples of asymmetriccycles are Wulff cracking processes and pressure swing reformingprocesses.

As mentioned above, regenerative reactors are known that separatelydeliver a stream of fuel, oxidant, or a supplemental amount of one ofthese reactants, directly to a location somewhere in the heat generationregion of the reactor. Although this may defer or control location ofcombustion, that process limits the cooling of the quench regions of thereactor, due to not having that stream pass through regenerative beds orregions. This can result in expanding heat zones loss of reactioncontrol.

The reactor heat generation region is typically a region of the reactorsystem that is in between two regenerative reactor beds or regions, withthe main regenerative flow passing from one of these bodies to theother. In most cases, this lateral stream is introduced via nozzles,distributors, or burners (e.g., Wulff) that penetrate the reactor systemusing a means that is generally perpendicular to flow direction andusually through the reactor vessel side wall. In large scale operations,such methods are impermissibly inefficient and costly. For example,during the exothermic step in a conventional Wulff cracking furnace, airflows axially through the regenerative bodies, and fuel is introducedvia nozzles that penetrate the side of the furnace, to combine with air(combusting and releasing heat) in an open zone between regenerativebodies. In a conventional symmetric RTO application, a burner is placedto provide supplemental combustion heat in a location in between tworegenerative bodies. The burner combusts fuel from outside the reactor,either with the air passing through the regenerative bodies, or usingexternal air. Additional measure must be made to ensure adequate andtimely quenching of the synthesized product, and to adequately cool thebed before the next cycle begins.

Attempts have been made to introduce a reactant of the exothermic stepto a location in the middle of the regenerative reactor via conduitsthat are positioned axially within one or more of the regenerativebodies. For example, Sederquist (U.S. Pat. No. 4,240,805) uses pipesthat are positioned axially within the regenerative bed to carry oxidant(air) to locations near the middle of the regenerative flow path.

All of these previously known systems suffer disadvantages that renderthe same inefficient and unpractical in any but very specialized, smallscale operations with methane feeds. Positioning nozzles, distributors,or burners in the middle of the regenerative flow path of the reactordiminishes the durability and control of the reactor system. Nozzles,distributors, and/or burners all rely on carefully-dimensioned passagesto regulate flow in a uniform manner, or to create the turbulence ormixing required to evenly distribute the heat that results from theexothermic reaction they support. These nozzles, distributors, and/orburners are located at the highest-temperature part of the reactor. Itis very difficult to fabricate and maintain carefully-dimensioned shapesfor use at high temperatures. If the nozzles or distributor loses itscarefully-dimensioned shape, it will no longer produce uniform flametemperatures.

A further disadvantage of separately or laterally introducing one ormore reactant directly into the middle or heat region of theregenerative flow path of the reactor is that such an arrangementbypasses that reactant around the regenerative flow path. In addition tonot quenching the quench portion of the reactor, such approach alsoeliminates preheating that reactant stream. The fundamental purpose of aregenerative reactor system is ideally to execute reactions at highefficiency by recuperating product heat directly into feeds. Bypassingsome fraction of the feed to the reactor around the regenerative systemthus reduces the efficiency potential of the reactor system and can leadto expanded heat zones and feed conversion reactions that last too long.

All of the known art disclose processes, methods, and equipment that areunsuitable for continuous, high efficiency operation at the necessarilyhigh temperature, due to complexity and thermal degradation ofequipment. Also, the known processes do not reliably provide methods ormeans for continuously controlling the location and dissipation of thecreated heat, resulting in either hot spots, undesired thermalmigration, and/or inefficient processes. What is needed is an efficientand cost-effective way to pyrolyze methane to acetylene at relativelyhigh yield, selectivity, and efficiency, in a manner that is competitivewith pyrolyzing other hydrocarbon feeds to acetylene.

SUMMARY OF THE INVENTION

The present inventors have discovered that acetylene can be efficientlymanufactured from methane feed according to the inventive reverse-flowregenerative reactor system, method, and process. This inventionprovides processes and apparatus for efficiently converting methane bycontrolling location, movement, and removal of reaction heat. Theinventive process beneficially feeds all of the exothermically reactingregeneration reactant streams through the recuperation or quenchingreactor bed media, while simultaneously deferring combustion, until thereactants reach a desired region of the reactor system. The inventionalso includes use of an inventive mixing apparatus within the high heatregion to provide efficient and complete mixing and exothermic reactionwithin the reactor system. The inventive process also preferablyutilizes hydrogen as a methane synthesis reaction diluent. The inventiveprocess creates and confines a regenerating “heat bubble” within thereactor system, without exposing any degradable components to the highheat. The inventive process consistently provides controlled exothermicreaction location and temperature migration and successfully avoidsequipment heat degradation.

In one preferred aspect, the invention includes a process for themanufacture of acetylene from methane feed using a cyclic reverse-flowreactor system, wherein the reactor system includes (i) a first reactorcomprising a first end and a second end, and (ii) a second reactorcomprising primary end and a secondary end, the first and secondreactors oriented in a series relationship with respect to each othersuch that the secondary end of the second reactor is proximate thesecond end of the first reactor, the process comprising the steps of:

-   -   (a) supplying a first reactant through a first channel in the        first reactor and supplying at least a second reactant through a        second channel in the first reactor, such that the first and        second reactants are supplied to the first reactor from the        first end of the first reactor;    -   (b) combining the first and second reactants at the second end        of the first reactor and reacting the combined reactants to        exothermically produce a heated reaction product;    -   (c) passing the heated reaction product through the second        reactor to transfer heat from the reaction product to the second        reactor;    -   (d) thereafter supplying methane through the heated second        reactor to the first reactor, to convert at least a portion of        the methane into acetylene;    -   (e) passing the supplied methane and the produced acetylene        through the first reactor to quench the methane and the produced        acetylene; and    -   (f) recovering the produced acetylene.

In another aspect, the invention includes a cyclic reverse flow reactorsystem for the manufacture of acetylene from methane feed, wherein thereactor system comprises:

(i) a first reactor comprising a first end and a second end;

(ii) a second reactor comprising primary end and a secondary end, thefirst and second reactors oriented in a series relationship with respectto each other such that the secondary end of the second reactor isproximate the second end of the first reactor;

wherein the first reactor further comprises;

(a) a first channel to supply at least a first reactant from the firstend of the first reactor to the second end of the first reactor;

(b) a second channel to supply at least a second reactant from the firstend of the first reactor to the second end of the first reactor; and

(c) a product removal line to remove at least one of methane and aproduced acetylene from the first reactor;

wherein the second reactor further comprises;

(i) a flue gas removal line to remove at least a portion of the heatedreaction product produced from mixing and reaction of the first andsecond reaction products; and

(ii) a methane feed line to feed methane to the primary end of thesecond reactor. The second reactor may be configured with a beddingarrangement similar to the first bed, or the second bed may beconfigured in separate fashion, such as according to a known pyrolysisbed design.

The first channel and the second channel in the first reactor maintainsseparated flow paths for the first and second reactants to prevent atleast a majority (by stoichiometric reactivity) of the first reactantand the second reactant from exothermically reacting with each otherwithin the first reactor. This defers reaction of the majority ofreactant until the same exits the second end of the first reactor.

BRIEF DESCRIPTION OF THE DRAWINGS

FIGS. 1( a) and 1(b) are a diagrammatic illustration of the two steps ina regenerating reverse flow reactor according to the present invention.

FIG. 2 is another diagrammatic illustration of an exemplary regenerativebed reactor that defers combustion, controls the location of theexothermic reaction, and adequately quenches the recuperation reactorbed.

FIG. 3 illustrates an axial view of an exemplary gas distributor.

FIG. 4 illustrates a cross sectional view of an exemplary gas mixer andchannels for controlled combustion. FIG. 4 a is a cutout view of aportion of FIG. 4.

FIG. 5 illustrates another exemplary embodiment of a reactor system,including a mixer and some illustrative component piping.

FIG. 6 illustrates yet another exemplary embodiment of a reactor systemthat utilizes separated, alternating layers of reactor bedding.

DETAILED DESCRIPTION

According to the invention, methane is converted to acetylene in areverse-flow reactor by pyrolysis, preferably at temperatures of fromabout 1500 to about 1900° C., and more preferably from about 1600 toabout 1700° C., with short residency times, e.g., less than 0.1 secondsand preferably less than about 5 milliseconds, and preferably in thepresence of hydrogen diluent. The conversion of methane into higherhydrocarbons such as acetylene requires a high reformation temperature,which in the past has been a barrier to commercialization andefficiency.

At least part of the invention of the present inventors is therecognition that the requisite high heat may be achieved by creating ahigh-temperature heat bubble in the middle of a packed bed system andthen use a two-step process wherein heat is (1) added to the bed viain-situ combustion, and then (2) removed from the bed via in-situendothermic reforming. A key benefit of the invention is the ability toconsistently manage and confine the high temperature bubble(e.g., >1600° C.) in a reactor region(s) that can tolerate suchconditions long term. The inventive process provides for a substantiallycontinuously operating, large-scale, cyclic, reverse-flow reactor systemthat is useful and operable on a commercial scale. This inventionovercomes the limitations of the prior art.

One common source for methane is natural gas. In some applications,natural gas, including associated hydrocarbon and impurity gases, may besupplied into the inventive reactor system. The supplied natural gasalso may be sweetened and/or dehydrated natural gas. Natural gascommonly includes various concentrations of associated gases, such asethane and other alkanes, preferably in lesser concentrations thanmethane. The supplied natural gas may include impurities, such as H2Sand nitrogen. The inventive methods and apparatus may also serve tosimultaneously convert some fraction of the associated higherhydrocarbons to acetylene. In other embodiments, the inventive methodsand compositions may be utilized with liquid feeds, such a vacuum gasoil (VGO) or naphthas.

The present invention may be described as methane pyrolysis in a reverseflow reactor system or more specifically the conversion of methane toacetylene via pyrolysis of methane in a reverse-flow reactor system. Thereactor system includes first and second reactors, oriented in a seriesrelationship with each other with respect to a common flow path, andpreferably along a common axis. The common axis may be horizontal,vertical, or otherwise. The present invention includes a processwherein: first and second in-situ combustion reactants are bothseparately, but preferably substantially simultaneously, passed througha quenching reactor bed (e.g., a first reactor bed), via substantiallyindependent flow paths (channels), to obtain the quenching (cooling)benefits of the total combined weight of the first and second reactants.(Although only first and second reactants are discussed, theregeneration reaction may also include additional reactants and reactantflow channels.) Both reactants are also concurrently heated by the hotquench bed, before they reach a designated location within the reactorsystem and react with each other in an exothermic reaction zone (e.g., acombustion zone). This deferred combustion of the first and secondreactants permits positioning initiation of the exothermic regenerationreaction at the desired location within the reactor system.

The reactants are permitted to combine or mix in the reaction zone tocombust therein, in-situ, and create a high temperature zone or heatbubble (1600-1700° C.) inside of the reactor system. Preferably thecombining is enhanced by a reactant mixer that mixes the reactants tofacilitate substantially complete combustion/reaction at the desiredlocation, with the mixer preferably located between the first and secondreactors. The combustion process takes place over a long enough durationthat the flow of first and second reactants through the first reactoralso serves to displace a substantial portion, (as desired) of the heatproduced by the reaction (e.g., the heat bubble), into and at leastpartially through the second reactor, but preferably not all of the waythrough the second reactor to avoid waste of heat and overheating thesecond reactor. The flue gas may be exhausted through the secondreactor, but preferably most of the heat is retained within the secondreactor. The amount of heat displaced into the second reactor during theregeneration step is also limited or determined by the desired exposuretime or space velocity that the methane feed gas will have to thereforming, high temperature second reactor media to convert the methaneand other hydrocarbons to acetylene.

After regeneration or heating the second reactor media, in thenext/reverse step or cycle, methane is supplied or flowed through thesecond reactor, from the direction opposite the direction of flow duringthe heating step. The methane contacts the hot second reactor and mixermedia, in the heat bubble region, to transfer the heat to the methanefor reaction energy. In addition to not wasting heat, substantiallyoverheating the reformer/second reactor bed may adversely lead to aprolonged reaction that cracks the methane past the acetylene generationpoint, breaking it down into its elemental components. Thus, the totalamount of heat added to the bed during the regeneration step should notexceed the sum of the heats that are required (a) to sustain thereforming reaction for the endothermic conversion of the suppliedmethane to acetylene for a suitable period of time, as determined bymany factors, such as reactor size, dimensions, gas flow rates,temperatures used, desired contact time, cycle duration, etc, and (b)for heat losses from the system both as conduction losses throughreactor walls as well as convective losses with the exiting products.The total amount of heat stored in the reactor system though isgenerally much more heat than would be minimally needed for conversionon any single cycle. However, it is desirable to avoid having thetemperature bubble so large that the residence time at temperaturebecomes too long. As is commonly done for reactor systems, normalexperimentation and refining adjustments and measurements can be made tothe reactor system to obtain the optimum set of reactor conditions.

In preferred embodiments, the reactor system may be described ascomprising two zones/reactors: (1) a heat recuperating (first)zone/reactor, and (2) a reforming (second) zone/reactor. As a catalystis not required to facilitate reforming methane to acetylene, in mostpreferred embodiments no catalyst is present in the reactor beds.However, there may be some applications that benefit from the presenceof a catalyst to achieve a certain range of reforming performance andsuch embodiments are within the scope of the invention.

The basic two-step asymmetric cycle of a reverse flow regenerative bedreactor system is depicted in FIGS. 1 a and 1 b in terms of a reactorsystem having two zones/reactors; a first or recuperator/quenching zone(7) and a second or reaction/reforming zone (1). Both the reaction zone(1) and the recuperator zone (7) contain regenerative beds. Regenerativebeds, as used herein, comprise materials that are effective in storingand transferring heat. The term regenerative reactor bed(s) means aregenerative bed that may also be used for carrying out a chemicalreaction. The regenerative beds may comprise bedding or packing materialsuch as glass or ceramic beads or spheres, metal beads or spheres,ceramic (including zirconia) or metal honeycomb materials, ceramictubes, extruded monoliths, and the like, provided they are competent tomaintain integrity, functionality, and withstand long term exposure totemperatures in excess of 1200° C., preferably in excess of 1500° C.,more preferably in excess of 1700° C., and even more preferably inexcess of 2000° C. for operating margin.

As shown in FIG. 1 a, at the beginning of the “reaction” step of thecycle, a secondary end (5) of the reaction zone (1) (a.k.a. herein asthe reformer or second reactor) is at an elevated temperature ascompared to the primary end (3) of the reaction bed (1), and at least aportion (including the first end (9)) of the recuperator or quench zone(7), is at a lower temperature than the reaction zone (1) to provide aquenching effect for the synthesis gas reaction product. A methanecontaining reactant feed, and preferably also a hydrogen diluent, isintroduced via a conduit(s) (15), into a primary end (3) of thereforming or reaction zone (1). Thereby, in a preferred embodiment, theterm pyrolysis includes hydropyrolysis.

The feed stream from inlet(s) (15) absorbs heat from the reformer bed(1) and endothermically reacts to produce the desired acetylene product.As this step proceeds, a shift in the temperature profile (2), asindicated by the arrow, is created based on the heat transfer propertiesof the system. When the bed is designed with adequate heat transfercapability, this profile has a relatively sharp temperature gradient,which gradient will move across the reaction zone (1) as the stepproceeds. The sharper the temperature gradient profile, the better thereaction may be controlled.

The methane/hydrogen/acetylene reaction gas exits the reaction zone (1)through a secondary end (5) at an elevated temperature and passesthrough the recuperator reactor (7), entering through a second end (11),and exiting at a first end (9) as a synthesized gas comprisingacetylene, some unconverted methane, and hydrogen. The recuperator (7)is initially at a lower temperature than the reaction zone (1). As thesynthesized reaction gas passes through the recuperator zone (7), thegas is quenched or cooled to a temperature approaching the temperatureof the recuperator zone substantially at the first end (9), which insome embodiments is preferably approximately the same temperature as theregeneration feed introduced via conduit (19) into the recuperator (7)during the second step of the cycle. As the reaction gas is cooled inthe recuperator zone (7), a temperature gradient (4) is created in thezone's regenerative bed(s) and moves across the recuperator zone (7)during this step. The quenching heats the recuperator (7), which must becooled again in the second step to later provide another quenchingservice and to prevent the size and location of the heat bubble fromgrowing progressively through the quench reactor (7). After quenching,the reaction gas exits the recuperator at (9) via conduit (17) and isprocessed for separation and recovery of the various components.

The second step of the cycle, referred to as the regeneration step, thenbegins with reintroduction of the first and second regenerationreactants via conduit(s) (19). The first and second reactants passseparately through hot recuperator (7) toward the second end (11) of therecuperator (7), where they are combined for exothermic reaction orcombustion in or near a central region (13) of the reactor system.

The regeneration step is illustrated in FIG. 1 b. Regeneration entailstransferring recovered sensible heat from the recuperator zone (7) tothe reaction zone (1) to thermally regenerate the reaction beds (1) forthe subsequent reaction cycle. Regeneration gas/reactants entersrecuperator zone (7) such as via conduit(s) (19), and flows through therecuperator zone (7) and into the reaction zone (1). In doing so, thetemperature gradients (6) and (8) may move across the beds asillustrated by the arrows on the exemplary graphs in FIG. 1( b), similarto but in opposite directions to the graphs of the temperature gradientsdeveloped during the reaction cycle in FIG. 1( a). Fuel and oxidantreactants may combust at a region proximate to the interface (13) of therecuperator zone (7) and the reaction zone (1). The heat recovered fromthe recuperator zone together with the heat of combustion is transferredto the reaction zone, thermally regenerating the regenerative reactionbeds (1) disposed therein.

In a preferred embodiment of the present invention, a first reactant,such as fuel, is directed down certain channels (each channel preferablycomprising a reactant flow path that includes multiple conduits) in thefirst reactor bed (7). In one embodiment, the channels include one ormore honeycomb monolith type structures. Honeycomb monoliths includeextruded porous structures as are generally known in the reactionindustry, such as in catalytic converters, etc. The term “honeycomb” isused broadly to refer to a cross-sectional shape that includes multipleflow paths or conduits through the extruded monolith and is not intendedto limit the structure or shape to any particular shape. The honeycombmonolith enables low pressure loss transference while providing contacttime and heat transfer. A mixer is preferably used between the zones toenable combustion. Each channel of the first and second channels isdefined broadly to mean the conductive conduit(s) or flow path(s) bywhich one of the reactants and synthesis gas flows through the firstreactor bed (7) and may include a single conduit or more preferably andmore likely, multiple conduits (e.g., tens, hundreds, or even thousandsof substantially parallel conduit tubes) that receive feed, such as froma gas distributor nozzle or dedicated reactant port.

The conduits each may have generally any cross-sectional shape, althougha generally circular or regular polygon cross-sectional shape may bepreferred. Each channel may preferably provide substantially parallel,generally common flow through the reactor media. Thus, a first channelmay be merely a single conduit, but more likely will be many conduits,(depending upon reactor size, flow rate, conduit size, etc.), forexample, such as exemplified in FIG. 2 and in Example 1 below. A channelpreferably includes multiple conduits that each receive and conduct areactant, such as delivered by a nozzle in a gas distributor. Theconduits may be isolated from each other in terms of cross flow alongthe flow path (e.g. not in fluid communication), or they may besubstantially isolated, such that reactant permeation through a conduitwall into the adjacent conduit is substantially inconsequential withrespect to reactant flow separation. One preferred reactor embodimentincludes multiple segments, whereby each segment includes a firstchannel and a second channel, such that after exiting the reactor, therespective first reactant is mixed with the respective second reactantin a related mixer segment. Multiple segments are included to providegood heat distribution across the full cross-sectional area of thereactor system.

Referring to FIG. 4, mixer segment (45), for example, may mix thereactant flows from multiple honeycomb monoliths arranged within aparticular segment. Each monolith preferably comprises a plurality (morethan one) of conduits. The collective group of conduits that transmitthe first reactant may be considered the first channel and a particularreactor segment may include multiple collective groups of monolithsand/or conduits conducting the first reactant, whereby the segmentcomprising a channel for the first reactant. Likewise, the secondreactant may also flow through one or more monoliths within a segment,collectively constituting a second channel. Thus, the term “channel” isused broadly to include the conduit(s) or collective group of conduitsthat conveys at least a first or second reactant. A reactor segment mayinclude only a first and second channel, or multiple channels formultiple flow paths for each of the first and second reactants. A mixersegment (45) may then collect the reactant gas from both or multiplechannels. Preferably, a mixer segment (45) will mix the effluent fromone first channel and one second channel.

It is recognized that in some preferred embodiments, several of theconduits within a channel will likely convey a mixture of first andsecond reactants, due at least in part to some mixing at the first end(17) of the first reactor. However, the numbers of conduits conveyingcombustible mixtures of first and second reactants is sufficiently lowsuch that the majority of the stoichiometrically reactable reactantswill not react until after exiting the second end of the first reactor.The axial location of initiation of combustion or exothermic reactionwithin those conduits conveying a mixture of reactants is controlled bya combination of temperature, time, and fluid dynamics. Fuel and oxygenusually require a temperature-dependent and mixture-dependentautoignition time to combust. Still though, some reaction will likelyoccur within an axial portion of the conduits conveying a mixture ofreactants. However, this reaction is acceptable because the number ofchannels having such reaction is sufficiently small that there is onlyan acceptable or inconsequential level of effect upon the overall heatbalance within the reactor. The design details of a particular reactorsystem should be designed so as to avoid mixing of reactants within theconduits as much as reasonably possible.

The process according to the present invention requires no largepressure swings to cycle the reactants and products through the reactorsystem. In some preferred embodiments, the reforming or pyrolysis ofmethane step occurs at relatively low pressure, such as less than about50 psia, while the regeneration step may also be performed at similarpressures, e.g., less than about 50 psia, or at slightly higher, butstill relatively low pressures, such as less than about 250 psia. Insome preferred embodiments, the methane pyrolysis step is performed at apressure of from about 5 psia to about 45 psia, preferably from about 15psia to about 35 psia. Ranges from about 5 psia to about 35 psia andfrom about 15 psia to about 45 psia are also contemplated. The mosteconomical range may be determined without more than routineexperimentation by one of ordinary skill in the art in possession of thepresent disclosure. Pressures higher or lower than that disclosed abovemay be used, although they may be less efficient. By way of example, ifcombustion air is obtained from extraction from a gas turbine, it may bepreferable for regeneration to be carried out at a pressure of, forexample, from about 100 psia to about 250 psia. However if, by way offurther example, the process is more economical with air obtained viafans or blowers, the regeneration may be carried out at 15-45 psia. Inone embodiment of the present invention, the pressure of the pyrolysisand regeneration steps are essentially the same, the difference betweenthe pressures of the two steps being less than about 10 psia.

It is understood that some method of flow control (e.g. valves, rotatingreactor beds, check valves, louvers, flow restrictors, timing systems,etc.) is used to control gas flow, actuation, timing, and to alternatephysical beds between the two flow systems. In the regeneration step,air and fuel must be moved through the reactor system and combined forcombustion. Air can be moved such as via compressor, blower, or fan,depending on the operating conditions and position desired for thereactor. If higher pressure air is used, it may be desirable to expandthe flue gas through an expansion turbine to recover mechanical energy.In addition, some fraction of exhaust gas may be recycled and mixed withthe incoming air. An exhaust gas recycle (EGR) stream may be suppliedwith at least one of the supplied first reactant and second reactant inthe first reactor. This EGR may be used to reduce the oxygen content ofthe regeneration feed, which can reduce the maximum adiabatic flametemperature of the regeneration feed. In the absence of EGR, CH4/airmixtures have a maximum adiabatic flame temperature of about 1980° C.;H2/air mixtures are about 2175° C. Thus, even if average temperature iscontrolled by limiting the flow rate of fuel, any poor diluting couldresult in local hot spots that approach the maximum flame temperature.Use of EGR can reduce the maximum hot spot temperature by effectivelyincreasing the amount of diluent such as N2 (and combustion products)that accompany the oxygen molecules.

For example, when 50% excess air is used for combustion, the maximumadiabatic flame temperature for H2-fuel/air combustion decreases fromabout 2175° C. to about 1640° C. Reducing the oxygen content of thesupplied air to about 13% would make about 1640° C. the maximumadiabatic flame temperature, regardless of local mixing effects. Thereforming or pyrolysis step and flow scheme is illustrated in FIG. 1(a). Methane (such as from natural gas) is supplied, preferably mixedwith or supplied with hydrogen as a diluent, either within the secondreactor or immediately prior to entry into the second reactor, and ispyrolyzed in the high temperature heat bubble created by theregeneration step. The methane containing feed may also includesubstantially any other hydrocarbon material that undergoes theendothermic reforming to acetylene, including natural gas, otherpetroleum alkanes, petroleum distillates, kerosene, jet fuel, fuel oil,heating oil, diesel fuel and gas oil, gasoline, and alcohols.Preferably, the feed will be gaseous material comprising methane and/orhydrocarbons that are in a gaseous state at the temperature and pressureof introduction into the reactor.

After leaving the second reactor and the optional mixer, the acetylenecontaining synthesized gas stream must be cooled or quenched to halt theconversion process at the acetylene stage. The timing for this step isimportant because acetylene is rarely a desired material for processexport. Rather, a preferred use for the produced acetylene is as anintermediate product in a flow process within a chemical plant, in routeto other preferred products, such as vinyl esters, ethylene,acetaldehyde, propanal, and/or or propanol, acrylic acid, and so on.After quenching, the synthesized gas stream may be provided to aseparation process that separates the acetylene, methane, hydrogen, andother gases. Recovered methane and hydrogen may be recycled forprocessing again in the reactor system. A separate process sequence mayconvert the acetylene to some other product. Each of these products maybe further processed to provide yet additional useful products, e.g.,acetaldehyde is typically an intermediate in the manufacture of ethanol,acetic acid, butanals, and/or butanols.

Ethylene is a basic building block of a plethora of plastics, and maytypically be the preferred use for the created acetylene, from theperspective of volume and value. Ethylene is conveniently manufacturedfrom acetylene by hydrogenation. In some embodiments of the invention,it may also be a coproduct of the inventive methane conversion process.

Another product of high interest is ethanol, which may be convenientlymanufactured by first hydrating the acetylene to acetaldehyde and thenhydrogenating acetaldehyde to ethanol. Ethanol is of interest because itis easily transported from a remote location and is easily dehydrated toethylene. Ethanol may also be suitable for use as a motor fuel, if themanufacturing can be sufficiently low in cost.

In any event, conversion of methane to acetylene leaves a surplus ofhydrogen. The idealized reaction to acetylene is:2CH₄→C₂H₂+3H₂ consuming about +45 kcal/mole of converted CH₄

As suggested by the above reaction, hydrogen is a valuable by-product ofthe present process. To a lesser extent, ethylene is also a valuableproduct, produced as a result of incomplete reduction of methane tohigher hydrocarbon. Unreacted methane is also a valuable product.

Accordingly, separation and recovery of hydrogen, separation andrecovery of ethylene, and separation and recovery of unconverted methaneare each individually and collectively preferred steps in the processaccording to the invention. Unconverted methane is preferably returnedto the hydropyrolysis reactor so that it may be converted on a secondpass. An amount of hydrogen should also be returned to thehydropyrolysis reactor that is sufficient to control the selectivity ofthe product distribution.

Since hydrogen is created (not consumed) in the reforming pyrolysisreaction, it will be necessary to purge hydrogen from the process in theamount of about one H2 for every CH4 converted. Hydrogen has a heat ofcombustion of about 57 Kcal/mole H2, so the hydrogen purged from theprocess has a heating value that is in the range of what is needed asregeneration fuel. Of course, if there is an alternate, high-value usefor the leftover hydrogen, then natural gas could be used for all orpart of the regeneration fuel. But the leftover hydrogen is likely to beavailable at low pressure and may possibly contain methane or otherdiluents. So, use of hydrogen as regeneration fuel may also be an idealdisposition in a remote location.

FIG. 2 illustrates another exemplary reactor system that may be suitablein some applications for controlling and deferring the combustion offuel and oxidant to achieve efficient regeneration heat. FIG. 2 depictsa single reactor system, operating in the regeneration cycle. Theinventive reactor system may be considered as comprising two reactorszones. The recuperator (27) is the zone primarily where quenching takesplace and provides substantially isolated flow paths or channels fortransferring both of the quenching reaction gases through the reactormedia, without incurring combustion until the gasses arrive proximate orwithin the reactor core (13) in FIG. 1. The reformer (2) is the reactorwhere regeneration heating and methane reformation primarily occurs, andmay be considered as the second reactor for purposes herein. Althoughthe first and second reactors in the reactor system are identified asseparately distinguishable reactors, it is understood and within thescope of the present invention that the first and second reactors may bemanufactured, provided, or otherwise combined into a common singlereactor bed, whereby the reactor system might be described as comprisingmerely a single reactor that integrates both cycles within the reactor.The terms “first reactor” and “second reactor” merely refer to therespective zones within the reactor system whereby each of theregeneration, reformation, quenching, etc., steps take place and do notrequire that separate components be utilized for the two reactors.However, most preferred embodiments will comprise a reactor systemwhereby the recuperator reactor includes conduits and channels asdescribed herein, and the reformer reactor may similarly possessconduits. Other preferred embodiments may include a reformer reactor bedthat is arranged different from and may even include different materialsfrom, the recuperator reactor bed. The bedding arrangement of thereformer or second reactor may be provided as desired or as prescribedby the application and no particular design is required herein of thereformer reactor, as to the performance of the inventive reactor system.Routine experimentation and knowledge of the methane pyrolysis art maybe used to determine an effective reformer/second reactor design.

As discussed previously, the first reactor or recuperator (27) includesvarious gas conduits (28) for separately channeling two or more gasesfollowing entry into a first end (29) of the recuperator (27) andthrough the regenerative bed(s) disposed therein. A first gas (30)enters a first end of a plurality of flow conduits (28). In addition toproviding a flow channel, the conduits (28) also comprise effective flowbarriers (e.g., which effectively function such as conduit walls) toprevent cross flow or mixing between the first and second reactants andmaintain a majority of the reactants effectively separated from eachother until mixing is permitted. As discussed previously, each of thefirst and second channels preferably comprises multiple channels or flowpaths. The first reactor may also comprise multiple substantiallyparallel flow segments, each comprising segregated first and secondchannels.

In a preferred embodiment of the present invention, the recuperator iscomprised of one or more extruded honeycomb monoliths. Preferredhoneycomb monoliths are extruded structures that comprise many (e.g., aplurality, meaning more than one) small gas flow passages or conduits,arranged in parallel fashion with thin walls in between. A small reactormay include a single monolith, while a larger reactor can include anumber of monoliths, while still larger reactor may be substantiallyfilled with an arrangement of many honeycomb monoliths. Each monolithmay be formed by extruding monolith blocks with shaped (e.g., square orhexagonal) cross-section and two- or three-dimensionally stacking suchblocks above, behind, and beside each other. Monoliths are attractive asreactor contents because they provide high heat transfer capacity withminimum pressure drop.

Each monolith may provide flow channel(s) (e.g., flow paths) for one ofthe first or second reactants. Each channel preferably includes aplurality of conduits. Alternatively, a monolith may comprise one ormore channels for each reactant with one or more channels or groups ofconduits dedicated to flowing one or more streams of a reactant, whilethe remaining portion of conduits flow one or more streams of the otherreactant. It is recognized that at the interface between channels, anumber of conduits will likely convey a mixture of first and secondreactant, but this number of conduits is proportionately small. In otherembodiments, a single flow channel may comprise multiple monoliths.Honeycomb monoliths preferred in the present invention (which areadjacent a first end (9) of the first reactor (7)) can be characterizedas having open frontal area (or geometric void volume) between about 40%and 80%, and having conduit density between about 50 and 2000 pores persquare inch, more preferably between about 100 and 1000 pores per squareinch. (For example, in one embodiment, the conduits may have a diameterof only a few millimeters, and preferably on the order of about onemillimeter.) Reactor media components, such as the monoliths oralternative bed media, preferably provide for at least one of the firstand second channels and preferably both channels to include a packingwith an average wetted surface area per unit volume that ranges fromabout 50 ft-1 to about 3000 ft-1, more preferably from about 100 ft-1 to2500 ft-1, and still more preferably from about 200 ft-1 to 2000 ft-1,based upon the volume of the first reactor that is used to convey areactant. These wetted area values apply to the channels for both of thefirst and second reactants. These relatively high surface area per unitvolume values are likely preferred for many embodiments to aid achievinga relatively quick change in the temperature through the reactor, suchas generally illustrated by the relatively steep slopes in the exemplarytemperature gradient profile graphs, such as in FIGS. 1( a), 1(b), and6. The quick temperature change is preferred to permit relatively quickand consistent quenching of the reaction to prevent the reaction fromcontinuing and creating coke.

Preferred reactor media components also provide for at least one of thefirst and second channels in the first reactor and more preferably forboth channels, to include a packing that includes a high volumetric heattransfer coefficient (e.g., greater than or equal to 0.02 cal/cm³s° C.,preferably greater than about 0.05 cal/cm³s° C., and most preferablygreater than 0.10 cal/cm³s° C.), have low resistance to flow (lowpressure drop), have operating temperature range consistent with thehighest temperatures encountered during regeneration, have highresistance to thermal shock, and have high bulk heat capacity (e.g., atleast about 0.10 cal/cm³° C., and preferably greater than about 0.20cal/cm³° C.). As with the high surface area values, these relativelyhigh volumetric heat transfer coefficient value and other properties arealso likely preferred for many embodiments to aid in achieving arelatively quick change in the temperature through the reactor, such asgenerally illustrated by the relatively steep slopes in the exemplarytemperature gradient profile graphs, such as in FIGS. 1( a), 1(b), and6. The quick temperature change is preferred to permit relatively quickand consistent quenching of the reaction to prevent the reaction fromcontinuing too long and creating coke or carbon buildup. The citedvalues are averages based upon the volume of reactor used for conveyanceof a reactant.

Alternative embodiments may use reactor media other than the describedand preferred honeycomb monoliths, such as whereby the channelconduits/flow paths may include a more tortuous pathways (e.g.convoluted, complex, winding and/or twisted but not linear or tubular),than the previously described monoliths, including but not limited tolabyrinthine, variegated flow paths, conduits, tubes, slots, and/or apore structure having channels through a portion(s) of the reactor andmay include barrier portion, such as along an outer surface of a segmentor within sub-segments, having substantially no effective permeabilityto gases, and/or other means suitable for preventing cross flow betweenthe reactant gases and maintaining the first and second reactant gasessubstantially separated from each other while axially transiting therecuperator (27). For such embodiments, the complex flow path may createa lengthened effective flow path, increased surface area, and improvedheat transfer. Such design may be preferred for reactor embodimentshaving a relatively short axial length through the reactor. Axiallylonger reactor lengths may experience increased pressure drops throughthe reactor. However for such embodiments, the porous and/or permeablemedia may include, for example, at least one of a packed bed, anarrangement of tiles, a permeable solid media, a substantiallyhoneycomb-type structure, a fibrous arrangement, and a mesh-type latticestructure. It may be preferred that the media matrix provides highsurface area to facilitate good heat exchange with the reactant andproduced gases.

It may be preferred to utilize some type of equipment or method toodirect a flow stream of one of the reactants into a selected portion ofthe conduits. In the exemplary embodiment of FIG. 2, a gas distributor(31) directs a second gas stream (32) to second gas stream channels thatare substantially isolated from or not in fluid communication with thefirst gas channels, here illustrated as channels (33). The result isthat at least a portion of gas stream (33) is kept separate from gasstream (30) during axial transit of the recuperator (27). In a preferredembodiment, the regenerative bed(s) of the recuperator zone comprisechannels having a gas or fluid barrier that isolates the first reactantchannels from the second reactant channels. Thereby, both of the atleast two reactant gases that transit the channel means may fullytransit the regenerative bed(s), to quench the regenerative bed, absorbheat into the reactant gases, before combining to react with each otherin the combustion zone.

As used in the present invention, gases (including fluids) (30) and (32)each comprise a component that reacts with a component in the otherreactant (30) and (32), to produce an exothermic reaction when combined.For example, each of the first and second reactant may comprise one of afuel gas and an oxidant gas that combust or burn when combined with theother of the fuel and oxidant. By keeping the reactants substantiallyseparated, the present invention defers or controls the location of theheat release that occurs due to exothermic reaction. By “substantiallyseparated” is meant that at least 50 percent, preferably at least 75percent, and more preferably at least 90 percent of the reactant havingthe smallest or limiting stoichiometrically reactable amount ofreactant, as between the first and second reactant streams, has notbecome consumed by reaction by the point at which these gases havecompleted their axial transit of the recuperator (27). In this manner,the majority of the first reactant (30) is kept isolated from themajority of the second reactant (32), and the majority of the heatrelease from the reaction of combining reactants (30) and (32) will nottake place until the reactants begin exiting the recuperator (27).Preferably the reactants are gases, but some reactants may comprise aliquid, mixture, or vapor phase.

The percent reaction for these regeneration streams is meant the percentof reaction that is possible based on the stoichiometry of the overallfeed. For example, if gas (30) comprised 100 volumes of air (80 volumesN2 and 20 Volumes O2), and gas (32) comprised 10 volumes of Hydrogen,then the maximum stoichiometric reaction would be the combustion of 10volumes of hydrogen (H2) with 5 volumes of Oxygen (O2) to make 10volumes of H2O. In this case, if 10 volumes of Hydrogen were actuallycombusted in the recuperator zone (27), this would represent 100%reaction of the regeneration stream. This is despite the presence ofresidual un-reacted oxygen, because that un-reacted oxygen was presentin amounts above the stoichiometric requirement. Thus, the hydrogen isthe stoichiometrically limiting component. Using this definition, it ispreferred than less than 50% reaction, more preferred than less than 25%reaction, and most preferred that less than 10% reaction of theregeneration streams occur during the axial transit of the recuperator(27).

In a preferred embodiment, the channels (28) and (33) comprise materialsthat provide adequate heat transfer capacity to create the temperatureprofiles (4) and (8) illustrated in FIG. 1 at the space velocityconditions of operation. Adequate heat transfer rate is characterized bya heat transfer parameter ATHT, below about 500° C., more preferablybelow about 100° C. and most preferably below about 50° C. The parameterATHT, as used herein, is the ratio of the bed-average volumetric heattransfer rate that is needed for recuperation, to the volumetric heattransfer coefficient of the bed, hv. The volumetric heat transfer rate(e.g. cal/cm³ sec) that is sufficient for recuperation is calculated asthe product of the gas flow rate (e.g. gm/sec) with the gas heatcapacity (e.g. ca./gm ° C.) and desired end-to-end temperature change(excluding any reaction, e.g. ° C.), and then this quantity divided bythe volume (e.g. cm³) of the recuperator zone (27) traversed by the gas.The ATHT in channel (28) is computed using gas (30), channel (33) withgas (32), and total recuperator zone (27) with total gas. The volumetricheat transfer coefficient of the bed, hv, is typically calculated as theproduct of a area-based coefficient (e.g. cal/cm2s° C.) and a specificsurface area for heat transfer (av, e.g. cm2/cm³), often referred to asthe wetted area of the packing.

In a preferred embodiment, channels (28) and (33) comprise ceramic(including zirconia), alumina, or other refractory material capable ofwithstanding temperatures exceeding 1200° C., more preferably 1500° C.,and still more preferably 1700° C. Materials having a workingtemperature of up to and in excess of 2000° C. might be preferred wherethere is concern with reaching the bed reaction adiabatic maximumtemperature for sustained periods of time, to prevent reactor beddamage, provided the project economics and conditions otherwise permituse of such materials. In a preferred embodiment, channels (28) and (33)have wetted area between 50 ft-1 and 3000 ft-1, more preferably between100 ft-1 and 2500 ft-1, and most preferably between 200 ft-1 and 2000ft-1. Most preferably, channel means (28) comprise a ceramic honeycomb,having channels running the axial length of the recuperator reactor(27).

Referring again briefly to FIGS. 1( a) and 1(b), the inventive reactorsystem includes a first reactor (7) containing a first end (9) and asecond end (11), and a second reactor (1) containing a primary end (3)and a secondary end (5). The embodiments illustrated in FIGS. 1( a),1(b), and 2 are merely simple illustrations provided for explanatorypurposes only and are not intended to represent a comprehensiveembodiment. Reference made to an “end” of a reactor merely refers to adistal portion of the reactor with respect to an axial mid-point of thereactor. Thus, to say that a gas enters or exits an “end” of thereactor, such as end (9), means merely that the gas may enter or exitsubstantially at any of the various points along an axis between therespective end face of the reactor and a mid-point of the reactor, butmore preferably closer to the end face than to the mid-point. Thereby,one or both of the first and second reactant gases could enter at therespective end face, while the other is supplied to that respective endof the reactor through slots or ports in the circumferential orperimeter out surface on the respective end of the reactor, such asillustrated in FIG. 6.

For example, in one embodiment, the channel segments could comprisehorizontal layers, (e.g., like a stack of pancakes) within a generallyrectangular box-like (not shown) or cylindrical-shaped first reactor,such as illustrated in FIG. 6, wherein each alternating layer conveysone of a first or second reactant while the two adjacent layers conveythe other reactant. The embodiment of FIG. 6 illustrates an exemplaryreactor system concept, wherein the first reactor comprises at least twolayers of reactor bed material and the first channel comprises at leastone of the at least two layers and the second channel comprises anotherof the at least two layers. Thus, every other layer may contain a feedport in the respective outer surface of the reactor, near the respectiveend of the reactor, to feed a reactant gas into the respective layers,while alternating layers are fed substantially simultaneously throughthe end face and/or by separate outer surface ports. A reactant flowpath through each of the at least two layers is substantially not influid communication with a flow path through an immediately adjacentlayer. A reactant flow path may preferably include one or moremonoliths, having a plurality of conduits, the aggregate of whichconduct one of the first or the second reactant through the firstreactor, without permitting a substantial amount of cross flow to andadjacent flow path. Some cross flow might be permitted to occur in someembodiments, provided the total amount of cross flow and resulting earlyreaction does not appreciably alter the acceptable heat balance in thesystem. Thereby, a majority of the stoichiometrically available firstand second reactants are available to react in the mixer (when present)and second reactor. FIG. 6 illustrates in simple form, how valves can beused to regulate flow through each of the layers of the first and secondreactors during each half of the full reaction cycle. In one set oflayers for supplying a common reactant, the side entry ports preferablyprovide entry for one reactant (e.g., air or fuel) that is isolated fromentry of the other reactant. Valves can control flow of one reactantinto (and out of) the side ports to alternate between “open” in theregeneration cycle and “closed” in the quenching cycle. When the firstreactor is flowing reactant into the side ports for regeneration, othervalves or lines leading from the respective ports/layers can close toprevent loss of the reactant. When the reactor is flowing generatedacetylene gas from the reactor, these latter side port valves can open,while the reactant supply valves close, to permit flow of the generatedacetylene-containing gas through all layers of the reactor. Although theillustrated exemplary embodiment of FIG. 6 includes side ports, thematerial entering through such ports is merely entering to gain accessto the first end of the first reactor and does not expose the inletcomponents to the heat bubble in the reactor core.

As an alternative to using only the side port valves to flow reactantfrom the side port layers, the first end (end-cap) of the reactor may beprovided with a set of flow restrictors on alternating layers, e.g.,such as single pane louvers, like a check-valve, or a double-pane“duck-bill” type of louver, that can act as a flow restrictor or valveto prohibit flow of the other reactant into the side port containinglayers from the end cap during regeneration, while permitting acetylenegas to flow through the side port layers and into the end cap plenumarea or out of the side ports while the reactor is quenching thegenerated acetylene. Flow restrictors may be positioned on an end face(9) to cover or seal over at least a portion of the conduits, such asevery other layer. The flow restrictors can close to permit flow intothe conduits of one or more of the layers from the end face, duringregeneration, and then open to permit flow through those layers duringreformation stage. During the regeneration stage, so long as thepressure on the end-face of the first reactor is higher than thepressure in the side port layers within the reactor, the side port gascan flow through the reactor without flowing out of the end-face, whilethe other reactant maintains the louvers/flow restrictor in closedposition to restrict or prevent flow of the other reactant into the sideport layers. The process may thereby include a step of actuating atleast one of a valve and/or a flow restrictor, such as a check valve, anactuated valve (e.g., electric, pneumatic, hydraulic, etc.) to preventthe undesirable flow of the restricted first or second reactant into theimproper layer or portion of the reactor. The restrictor may open orclose in response to pressure and pressure changes within the firstportion/layers and second portion/layers of the reactor. Preferably, theflow restrictor(s) is passively responsive to pressure changes withinthe system, such as use of a hinge system. Thereby, the first reactorcan utilize substantially the whole reactor bed and both of the firstand second reactants to cool the quenching reactor and for heating thefirst and second reactants prior to their combining and reacting.

With regard to the various exemplified embodiments, FIG. 3 illustratesan axial view of an exemplary gas distributor (31) having apertures(36). Referring to both FIGS. 2 and 3, apertures (36) may direct thesecond reactant gas (32) preferentially to select channels (33). In apreferred embodiment, apertures (36) are aligned with, but are notsealed to, the openings/apertures of select channels (33). Nozzles orinjectors (not shown) may be added to the apertures (36) that aresuitably designed to direct the flow of the second gas (32)preferentially into the select channels (33). By not “sealing” the gasdistributor apertures (36) (or nozzles/injectors) to the select channels(33), these channels may be utilized during the reverse flow or reactioncycle, increasing the overall efficiency of the system. Such “open” gasdistributor (31) may be preferred for many applications, over a “closed”system, to facilitate adaptation to multiple reactor systems, such aswhere the reactor/recuperator beds may rotate or otherwise move inrelation to the location of the gas stream for processing, e.g., such aswith a rotating bed type reactor system.

When a gas distributor nozzle or aperture (36) in an “open” systemdirects a stream of reactant gas (32) toward the associated inletchannel and associated conduits in the reactor (preferably a honeycombmonolith(s)), the contents of that stream of reactant gas (32) willtypically occupy a large number of honeycomb conduits (33) as ittraverses the recuperator. This outcome is a geometric result of thesize of the reactor segments and/or aperture size, relative to the sizeof the monolith honeycomb conduits. The honeycomb conduits occupied bygas (32) may, according to a preferred embodiment, be characterized as abundle of conduits, typically oriented along the same axis as theaperture (36) and its issuing stream of gas (32). Conduits located nearthe center of this bundle/channel will contain a high purity of gas (32)and thus will likely not support exothermic reaction. Conduits locatednear the outer edge of the bundle will be in close proximity to conduits(28) carrying the other reactant. In an “open” system as describedabove, some mixing of the first gas (30) and the second gas (32) will beunavoidable near the peripheral edges of each stream of gas (32) thatissues from the apertures (36). Thus, some conduits (28) and (33) nearthe outer edge of the bundle will carry some amount of both the firstgas (30) and the second gas (32). Reaction or combustion between gases(30) and (32) could happen in these conduits before the gases completelytraverse recuperator (27). Such gases would still be considered to besubstantially separated, as long as the resulting reaction of theregeneration streams within the recuperator (27) is less than 50%,preferably than less than 25%, and most preferably less than 10% of thestoichiometrically reactive reactant having the smallest or reactionlimiting presence.

In some alternative embodiments, the recuperator reactor (27) mayinclude, for example, packed bed or foam monolith materials (not shown)that permit more mixing or dispersion of reactants before fullytraversing the first reactor. In this case, additional reaction mayoccur in the recuperator (27) due to mixing within the recuperator thatis due to the axial dispersion of gases (30) and (32) as they passthough. This may still be an acceptable arrangement as long as themixing and subsequent reaction of the regeneration streams within therecuperator (27) is less than 50%, preferably than less than 25%, andmost preferably less than 10%. Methods for calculation of radialdispersion and mixing in bed media is known in the art.

During regeneration, the first gas (30) and second gas (32) transit therecuperator zone (27) via channels (28) and (33). It is a key aspect ofthis invention that heat, stored in the recuperator zone from theprevious quench cycle, is transferred to both the first and second gasesduring the regeneration cycle. The heated gases are then introduced intomixer (44). The gas mixer (44), located between the recuperator (27) andthe reactor (21), functions to mix the regenerating reaction gas streams(30) and (32), preferably at or near the interface of the reaction zone(21) and the mixer (44).

The mixer (44) is preferably constructed or fabricated of a materialable to withstand the high temperatures expected to be experienced inthe reaction zone during methane reforming at high selectivity and highconversion rates (>50 wt %). In a preferred embodiment, mixer (44) isconstructed from a material able to withstand temperatures exceeding1200° C., more preferably 1500° C., and most preferably 1700° C. In apreferred embodiment, mixer means (34) is constructed of ceramicmaterial(s) such as alumina or silicon carbide for example.

FIG. 4 illustrates an axial view of one configuration of the mixer (44),together with a cut-away view FIG. 4 a, of one exemplary embodiment ofswirl-type mixer (47). The exemplary mixer (44) comprises mixer segments(45) having swirl mixer (47) located within the sections (45). In apreferred embodiment, mixer segments (45) are substantially equal incross sectional area and the swirl mixers (47) are generally centrallylocated within the sections (45). Mixer segments (45) are positionedwith respect to the reactor system to segment the gas flow of aplurality of gas channels (28) and (33). In a preferred embodiment,segments (45) may each have substantially equal cross sectional area tofacilitate intercepting gas flow from a substantially equal number ofgas channel means (28) and (33). Also in a preferred embodiment, the gaschannels (28) and (33) are distributed within recuperator reactor (27)such that each of the segments (45) intercepts gas flow from asubstantially equal fraction of both first gas channel means (28) andsecond gas channel means (33). Expressed mathematically, one can definefAi as the fraction of total cross sectional area encompassed by sectioni, f28 i as the fraction of total channel means (28) intercepted bysection i, and f33 i as the fraction of total channel means (33)intercepted by section i. In a preferred embodiment, for each section i,the values f28 i, and f33 i will be within about 20% of (i.e. betweenabout 0.8 and 1.2 times) the value of fAi, and more preferably withinabout 10%. One can further define f30 i as the fraction of gas stream(30) intercepted by section i, and f32 i as the fraction of gas stream(32) intercepted by the section i. In a more preferred embodiment, foreach section i, the values of f30 i, and f32 i will be within about 20%of fAi, and more preferably within about 10%.

FIG. 4 a illustrates an exemplary cut out section of an individual gasmixer segment (45) with swirl mixer (47). While the present inventionmay utilize a gas mixer known to the skilled artisan to combine gasesfrom the plurality of gas channel means (28) and (33), a preferredembodiment of this invention minimizes open volume of the gas mixer (44)while maintaining sufficient mixing and distribution of the mixed gases.The term open volume means the total volume of the swirl mixers (47) andgas mixer segment (45), less the volume of the material structure of thegas mixer. Accordingly, gas mixer segment (45) and swirl mixer (47) arepreferably configured to minimize open volume while concurrentlyfunctioning to provide substantial gas mixing of the gases exiting gaschannels (28) and (33). In a preferred practice of the invention, gasmixer segment (45) dimensions L and D, are tailored to achievesufficient mixing and distribution of gases (31) and (32) whileminimizing open volume. Dimension ratio L/D is preferably in the rangeof 0.1 to 5.0, and more preferably in the range of 0.3 to 2.5. Forgeneral segments of area A, a characteristic diameter D can be computedas 2(A/π)½.

In addition, the total volume attributable to the gas mixer (44) ispreferably tailored relative to the total volume of the first reactorbed (27) and reforming bed (21). Gas mixer (44) preferably has a totalvolume of less than about 20%, and more preferably less than 10% of thecombined volume of the recuperator zone (27), the reformation zone (21),and the gas mixer (44).

Referring again to FIG. 2, the mixer (44) as configured combines gasesfrom channels (33) and (28), and redistributes the combined gas acrossand into reaction zone (21). In a preferred embodiment, first reactantand second reactant are each a gas and one comprises a fuel and theother an oxidant. Fuel may comprise hydrogen, carbon monoxide,hydrocarbons, oxygenates, petrochemical streams, or mixtures thereof.Oxidant typically comprises a gas containing oxygen, commonly mixed withN2, such as air. Upon mixing, the fuel and oxidant at mixer (44), thegases combust, with a substantial proportion of the combustion occurringproximate to the entrance to the reaction zone (21).

The combustion of the fuel and oxygen-containing gas proximate to theentrance of the reformer or reaction zone (21) creates a hot flue gasthat heats (or re-heats) the reaction zone (21) as the flue gas travelsacross that zone. The composition of the oxygen-containing gas/fuelmixture is adjusted to provide the desired temperature of the reactionzone. The composition and hence reaction temperature may be controlledby adjusting the proportion of combustible to non-combustible componentsin the mixture. For example, non-combustible gases or other fluids suchas H2O, CO2, and N2 also may be added to the reactant mixture to reducecombustion temperature. In one preferred embodiment, non-combustiblegases comprise steam, flue gas, or oxygen-depleted air as at least onecomponent of the mixture.

Referring again to regeneration FIG. 1( b), the reacted, hot combustionproduct passes through reformer (1), from the secondary end (5) to theprimary end (3), before being exhausted via conduit (18). The flow ofcombustion product establishes a temperature gradient, such asillustrated generally by example graph (8), within the reformation zone,which gradient moves axially through the reformation reaction zone. Atthe beginning of the regeneration step, this outlet temperature maypreferably have an initial value substantially equal (typically within25° C.) to the inlet temperature of the reforming feed of the preceding,reforming, step. As the regeneration step proceeds, this outlettemperature will increase somewhat as the temperature profile movestoward the outlet, and may end up 50 to 200° C. above the initial outlettemperature.

The inventive processes and apparatus has been described generally andwith regard to illustrative embodiments, such as provided within theaccompanying figures. The invention has overcome the limitations thatrendered the prior art impractical and non-useful for practicing methanereforming to acetylene. The following text elaborates on the previousdiscussion and discloses some preferred embodiments of the methods andapparatus for practicing the invention.

Referring again to FIGS. 1( a) and 1(b), the invention includes apreferred process for the manufacture of acetylene from methane feedusing a reverse-flow reactor system, wherein the reactor system includes(i) a first reactor (7) comprising a first end (9) and a second end(11), and (ii) a second reactor (1) comprising primary end (3) and asecondary end (5), the first and second reactors oriented in a seriesrelationship with respect to each other such that the secondary end ofthe second reactor is proximate the second end of the first reactor, theprocess comprising the steps of: (a) supplying a first reactant througha first channel in the first reactor and supplying at least a secondreactant through a second channel in the first reactor, such that thefirst and second reactants are supplied, such as via conduit (19) to thefirst reactor from the first end of the first reactor; (b) combining thefirst and second reactants at the second end of the first reactor andreacting the combined reactants to exothermically produce a heatedreaction product; (c) passing the heated reaction product through thesecond reactor (1) to transfer heat from the reaction product to thesecond reactor; (d) thereafter supplying methane through the heatedsecond reactor to the first reactor, such as via conduit (15) to convertat least a portion of the methane into acetylene; (e) passing thesupplied methane and the produced acetylene through the first reactor toquench the methane and the produced acetylene; and (f) recovering theproduced acetylene, such as via conduit (17). The first and secondreactors are oriented in a series relationship with respect to eachother means that the reactors are in series with respect to a commonflow path, such that material exiting from one of the reactors flowsinto the other reactor, regardless of direction of flow. There may alsobe space or components intermediate the first and second reactors, suchas a mixer. The secondary end of the second reactor is proximate thesecond end of the first reactor is provided merely for orientationreference. Flow of material exiting the second end of the first reactorfollows the flow path to enter the secondary end of the second reactor,and vice versa. Preferably, the flow path through the reactors share acommon axis but the reactors may be arranged otherwise, such as besideeach other with the flow path generally forming a “U” shape. Also, thespacing between the reactors is subject to adjustment as determined bythe presence of mixers or other apparatus, but it may be preferred thatthey are closer together as opposed to more distant between them, suchthat the reaction products may move quickly between the reactors.

Preferably, the cyclic reverse flow reactor system further comprises amixer (not shown in FIG. 1, but illustrated in FIG. 2 as component (44))situated intermediate (13) the first reactor and second reactor to mixthe first reactant with the second reactant and more preferably, thefirst channel and the second channel axially traverse the first reactorto pass the first and second reactants to the mixer; and further, thatthe first and second reactants are combined at the second end of thefirst reactor such that less than half of the total weight of thecombined first and second reactant supplied, based upon the total weightof the first and second reactant, is reacted exothermically before thesupplied first and second reactants exit the second end of the firstreactor.

The preferred process may also include the optional step of supplyinghydrogen in the second reactor to moderate the reaction of the methane.The supplied hydrogen may be as hydrogen gas, a mixture of gasescomprising hydrogen, steam, or another product that provides or yieldshydrogen in the reactor. The hydrogen may be mixed with the methane sothat the mixture is introduced into the reforming reactor, or each ofthe methane and hydrogen gas may separately be introduced into thereformer. In a preferred embodiment, hydrogen is mixed with the methanebefore introduction into the reactor. Preferred molar ratio of hydrogento methane is between 0 and 5. More preferred ratio of hydrogen tomethane in the present invention is between about 1 and about 3.

The first and second regeneration reactants may exothermically react inat least one of (i) the second end of the first reactor, such as atleast partially within a second end of the recuperator, (ii) a regionintermediate the first and second reactors, such as within a mixer orspace between the first and second reactors, and (iii) the secondary endof the second reactor, such as at least partially within a secondary endof the reformer reactor, including at an interface zone between areformer side of the mixer and the reformer reactor. Preferably, theregion (13) intermediate the first and second reactors includes a mixerfor mixing and combusting at least a portion of the first and secondreactants.

It may also be preferred in some embodiments that, in step (b) above,the first and second reactants are combined in a mixer (44) positionedproximate the second end of the first reactor, wherein at least about75% of the regeneration stream is reacted. Thus, up to about 25 wt % ofthe combined regeneration stream may have combusted or reactedsubstantially within the first reactor, preferably nearer the second endof the first reactor than the first end. This leaves about 75 wt % toreact to provide heat to the mixer and reformer/second reactor (1). Insome alternative embodiments, up to 50 weight percent of the combinedreactants may be permitted to react before leaving the first reactor. Ina most preferred embodiment, 90% of the regeneration stream is reactedin the mixer or in a region of reformer (21) proximate to the mixer(near the second end, near end (5) in FIG. 1 b).

For those embodiments comprising a mixer, the mixer should beconstructed from material able to withstand temperatures in excess ofabout 1200° C., preferable to withstand temperatures in excess of about1500° C., and most preferably to withstand temperatures in excess ofabout 1700° C. In some preferred embodiments, the mixer comprises arefractory material, such as a ceramic. In a preferred embodiment, thefirst reactor or regenerator (7) has a geometric void volume A, and thesecond reactor (1), has a geometric void volume B, and the mixer (44)(including void volume (13) for the embodiments not having a mixerdevice but instead merely providing a region for the gases to comingle,mix, and combust) has a geometric void volume C, whereby void volume Cis less than or equal to about 20 percent and preferably less than orequal to about 10 percent, of the total combined void volumes A plus Bplus C. The term geometric void is used herein to denote the void volumein major passages that gases use to transit the reactor, and to excludevolumes than may be present in small pores within the walls of thereactor contents. For example, for honeycomb monoliths, geometric volumeincludes the volume in the channels, but excludes any pore volume thatmay exist in the channel walls.

Preferably, the first reactant used in the exothermic regenerativereaction includes a fuel comprising CO, H2, hydrocarbon(s), oxygenates,petrochemicals, or a mixture thereof. Other components, reactive and/ornon-reactive, may also be present in the fuel. The second reactant maycomprise oxygen, such as from air. In the reforming step, the suppliedmethane may also be supplied from the primary end of the second reactorto act as a diluent or to react with the methane in the second reactor.

According to some preferred processes, the step of supplying the methanefeed through the second reactor is performed at a pressure in the secondreactor of from about 5 psia up to about 45 psia, or more preferably ata pressure in the second reactor of from about 15 psia up to about 35psia. The regeneration step of supplying at least one of the firstreactant and the second reactant to the first reactor may be performedat a relatively low pressure, such as about the pressure that the stepof reforming is performed. Alternatively, such as in an embodiment wherethe regeneration step is supplied a regeneration reactant including fueland/or the exhaust gas from a turbine, compressor, or blower, theregeneration step may be performed at a pressure great than about 35psia, such as a pressure of preferably up to about 250 psia. A largerpressure may be utilized. However equipment and process considerationsmay make use of such pressures undesirable. Preferably, the heatedreaction product in the regeneration step heats at least a portion ofthe second reactor, preferably the secondary end (5) of the secondreactor (1), to a temperature of at least about 1500° C., and morepreferably to a temperature of at least about 1600° C.

Reactor system cycle time includes the time spent at regeneration plusthe time spent at reforming, plus the time required to switch betweenregeneration and reformation and vice versa. Thus, a half cycle may bethe substantially the time spent only on regeneration, or the time spendon reformation. A complete cycle includes heating the bed, feeding themethane, and quenching the acetylene containing reaction product.Typical cycle times for preferred embodiments utilizing honeycombmonoliths may be between 1 second and 240 seconds, although longer timesmay be desired in some alternative embodiments. More preferably for thepreferred monolith embodiments, cycle times may be between 2 seconds and60 seconds. It is not necessary that the regeneration and reformationsteps to have equal times, and in a well-refined application it islikely that these two times will not be equal.

Also, although not required for reforming methane to acetylene, in somealternative embodiments, the reforming/second reactor (1) may furthercomprise a reaction catalyst. After the methane has been reformed andpassed through the quenching first reactor (7), the process may alsoinclude the step of recovering the acetylene from the quenchedacetylene-methane mixture. Such recovery processes may also include thestep(s) of recovering at least one of hydrogen and methane from thequenched acetylene-methane mixture for recycling to the second reactor.

In a broad aspect, the inventive process includes a process for thepyrolysis manufacture of acetylene, comprising the steps of: (a)supplying a first reactant through a first portion of a reactor bed; (b)supplying at least a second reactant through a second portion of thereactor bed substantially separate from the first portion of the reactorbed, and (c) combining the supplied first reactant with the secondreactant after the first and second reactants have separately traversedat least a portion of the reactor bed, for the first and secondreactants to exothermically react with each other; wherein both of thefirst and second portions of the reactor bed are utilized to quench asynthesized reaction product comprising acetylene, after the combinedfirst and second reactants have reacted with each other. Preferably thesynthesized reaction product is a product of pyrolysis of a feed thatincludes methane. According to other embodiments, the synthesizedreaction product is a product of pyrolysis of a feed that includeshydrocarbons other than methane, such as ethane, propane, naphtha, orother pyrolyzable hydrocarbons, and/or feeds that include methane as acomponent therein. Although the present invention pertains primarily toinventive processes, methods, and equipment for converting methane toacetylene, the subject processes and equipment may also be utilized forpyrolysis of feeds other than methane, including liquid and/or gasfeeds, into any of a number of desirable pyrolysis reaction products.

The invention also includes the reactor system equipment and apparatusutilized in performing the inventive processes. According to onepreferred embodiment, the subject invention includes a cyclic reverseflow reactor for the manufacture of acetylene from methane feed, whereinthe reactor includes a reactor system that comprises: (i) a firstreactor comprising a first end and a second end; (ii) a second reactorcomprising primary end and a secondary end, the first and secondreactors oriented in a series relationship with respect to each othersuch that the secondary end of the second reactor is proximate thesecond end of the first reactor; wherein the first reactor furthercomprises; (a) a first channel to supply at least a first reactant fromthe first end of the first reactor to the second end of the firstreactor; (b) a second channel to supply at least a second reactant fromthe first end of the first reactor to the second end of the secondreactor; and (c) a product removal line to remove at least one ofmethane and a produced acetylene from the first reactor; wherein thesecond reactor further comprises; (d) a flue gas removal line to removeat least a portion of the heated reaction product produced from mixingand reaction of the first and second reaction products; and (e) amethane feed line to feed methane to the primary end of the secondreactor. Preferably, at least one of the first channel and the secondchannel prevent a stoichiometric reactable majority of the reactionlimiting first reaction product and second reaction product, fromexothermically reacting with each other to produce a reaction productuntil the unreacted first and second reaction products exit the secondend of the first reactor, based upon the total combined weight of thefirst and second reaction products. Preferably, the reactor furthercomprises a mixer situated intermediate the first reactor and secondreactor to mix the first reactant with the second reactant. Thepreferred mixer may also further comprise one or more gas mixersegments, wherein each segment receives at least a portion of the firstand second reactant to mix the at least a portion of the first andsecond reactant in the respective segment. In preferred embodiments, themixer is constructed from material able to withstand temperatures inexcess of about 1200° C., more preferably in excess of about 1500° C.,still more preferably in excess of about 1700° C., and yet still morepreferably in excess of about 2000° C. According to some embodiments,the mixer comprises a ceramic. Preferably the first reactor has a voidvolume A, and the second reactor has a void volume B, and the mixer hasa void volume C, whereby void volume C is less than or equal to aboutfifty percent and more preferably less than or equal to about twentypercent of the total of void volume A plus void volume B plus volume C.

The reactor further comprises a methane/feed supply line to supplymethane, and or any optional other gases, such ashydrogen/diluent/additional reactants, to the primary end of the secondreactor for conversion to acetylene in the second reactor. The reactormay also include hydrogen diluent or reaction component (including acomponent that comprises hydrogen) supply line to supply hydrogen to theprimary end of the second reactor so that the hydrogen can react withthe methane. The reactor also includes a fuel gas supply line to supplya fuel gas to one of the first channel and the second channel in thefirst reactor, and a second reactant supply line to supply a secondreactant to react with the first reactant. The second reactantpreferably comprises oxygen and the second reactant may include air. Thesecond reactant may also include at least one of (i) a noncombustiblegas, and (ii) a mixture of combustible and noncombustible gases, such asan exhaust gas recycle (EGR).

Example 1

The following example is merely illustrative of one exemplary embodimentand process, and is not intended to limit the scope of the invention.Methane is converted to acetylene in a pair of reactor systems, eacharranged according to simplified illustration FIG. 5, suitably valved,such that one reactor system is executing the regeneration step whilethe other reactor system is executing the pyrolysis step. Someembodiments may also include more reactor systems than just a pair, suchas multiple reactor systems, each operating in a phased timingarrangement, such that the entire process is substantially continuous.The reactor system includes at least a first reactor/recuperator (102),a mixer or mixing zone (109), and a second reactor/reforming zone (101).

Both reforming (101) and recuperation (102) reactor zones compriseextruded, ceramic honeycomb monolith blocks, stacked in 3 dimensions tofill the reactor zone volume. The overall reactor systems are about 10ft in diameter and about 4 ft in height or axial length (excludinginsulation and vessel shell that surrounds the reactor). Recuperationzone (102) is at the top, and measures about 10 ft in diameter by about17 inches long. Mixer zone (109) is about 10 ft in diameter by about 4inches long and reforming zone (101) is about 10 ft in diameter by about27 inches long. The ceramic honeycomb monoliths have conduit structurecharacterized as having about 400 conduits or pores per square inch and56% geometric void volume. The mixer zone is comprised of manyindividual mixer blocks, each block representing a mixer section (45) asdescribed with respect to FIG. 4. Blocks are hexagonal in cross-sectionand measure 3 inches across the flats of the hexagon and 4 inches inheight, and function to mix the regeneration fuel and oxidant reactants,and redistribute flow to the conduits in downstream zone. Individualmixer blocks are stacked side by side to fill the 10-ft diameter mixerzone (109).

During the regeneration step, regeneration oxidant (106) is fed intoplenum (107) located above the recuperation zone (102). Regenerationfuel (105) is fed into a fuel distributing sparger or distributor (110)that is located within the plenum and which has one fuel orifice (suchas aperture (36) in FIG. 3) positioned axially above each of the mixerblock segments in the mixer zone (109), whereby each orifice feeds fuelinto a group of conduits thus forming a fuel channel. Combustionproduct, or flue gas (111) resulting from the combustion of theregeneration fuel (105) and oxidant (106) is withdrawn from the plenum(108) that is below the reforming zone and that accesses the individualconduits in the reforming zone.

During the pyrolysis step, pyrolysis feed (103) is introduced into thereforming zone (101) by way of the manifold or plenum (108). If desired,a sparger or distributor (not shown) may also be used within the plenum(108) to supply the methane feed into the reformer (101). Pyrolysis feedis converted to acetylene as it travels through the reforming zone(101), the mixer zone (109) and the recuperator zone (102). Quenchedpyrolysis product (104) is recovered from the plenum (107) above therecuperator zone (102).

The regeneration step and pyrolysis step are each operated for aselected duration, which for this Example is three seconds in eachdirection, before switching to the alternate step, such that thecomplete cycle requires about six seconds plus switching time.Composition and flow rate of the total streams are shown in Table 1,below, for the pair of reactor systems. These flows representsubstantially continuous flows, as one reactor is always in pyrolysiswhile the other reactor is in regeneration stage. This example issomewhat idealized, because it provides no time to switch the reactorfrom one step to the next. More sophisticated cycles that provide forsuch switching may be added by one skilled in the art.

TABLE 1 Stream 103 104 105 106 111 T, ° C. 100 354 100 100 265 kg/hrkg/hr kg/hr kg/hr kg/hr C3+ 476 C2H2 21,740 C2H4 2,162 C2H6 81 CH447,032 14,795 2,436 H2O 29,485 H2 11,696 17,670 2,630 CO2 13,166 N2229,028 229,042 O2 65,470 29,686 Total 58,728 56,924 5,066 294,498301,379

This Example demonstrates the successful conversion of methane toacetylene and hydrogen, and supports the following conclusions: (1) Highproductivity of acetylene. The inventive reactor system can produceacetylene yields that are many times (e.g., >10 times) higher than priorart Wulff-type reactors. (2) The process and reactor system succeedswithout use of metals or degradable components exposed to the hot zone.(3) High selectivity and yield is achieved on a commercial size reactorsystem, demonstrating that the inventive system is useful for highproduction rates, such as on a commercial scale.

Trade names used herein are indicated by a ™ symbol or ® symbol,indicating that the names may be protected by certain trademark rights,e.g., they may be registered trademarks in various jurisdictions. Allpatents and patent applications, test procedures (such as ASTM methods,UL methods, and the like), and other documents cited herein are fullyincorporated by reference to the extent such disclosure is notinconsistent with this invention and for all jurisdictions in which suchincorporation is permitted.

When numerical lower limits and numerical upper limits are listedherein, ranges from any lower limit to any upper limit are contemplated.While the illustrative embodiments of the invention have been describedwith particularity, it will be understood that various othermodifications will be apparent to and can be readily made by thoseskilled in the art without departing from the spirit and scope of theinvention. Accordingly, it is not intended that the scope of the claimsappended hereto be limited to the examples and descriptions set forthherein but rather that the claims be construed as encompassing all thefeatures of patentable novelty which reside in the present invention,including all features which would be treated as equivalents thereof bythose skilled in the art to which the invention pertains.

The invention has been described above with reference to numerousembodiments and specific examples. Many variations will suggestthemselves to those skilled in this art in light of the above detaileddescription. All such obvious variations are within the full intendedscope of the appended claims.

What is claimed is:
 1. A process for the pyrolysis manufacture ofacetylene, comprising the steps of: (a) supplying a first reactantthrough a first portion of a reactor bed; (b) supplying at least asecond reactant through a second portion of the reactor bedsubstantially separate from the first portion of the reactor bed; and(c) combining the supplied first reactant with the second reactant afterthe first and second reactants have separately traversed at least aportion of the reactor bed, for the first and second reactants toexothermically react with each other; wherein both of the first andsecond portions of the reactor bed are utilized to quench a pyrolysisreaction product comprising acetylene, after the combined first andsecond reactants have reacted with each other.
 2. The process of claim1, wherein the reaction product is a product of pyrolysis of a feed thatincludes methane.
 3. The process of claim 1, wherein the synthesizedreaction product is a product of pyrolysis of a feed that includeshydrocarbons other than methane.
 4. The process of claim 1, wherein thefirst reactant is supplied through the reactor bed substantiallysimultaneously with the supplying of the at least a second reactantthrough the reactor bed.